Method for two-step hydrocracking of hydrocarbon feedstocks

ABSTRACT

An improved hydrocracking process of hydrocarbon charges, in two stages with an intermediate separation, in which the second-stage of hydrocracking is carried out in the presence of an added nitrogen content which is greater than 150 ppm by weight, preferably about 500 ppm by weight, and preferably in the presence of a Y zeolite catalyst, produces high yields of middle distillate.

The present invention relates to an improved hydrocracking process ofhydrocarbon charges, a process in two-stages with intermediateseparation, in which the second-stage of hydrocracking is carried out inthe presence of an added nitrogen content which is greater than 110 ppmby weight.

The aim of the process is essentially the production of middledistillates, that is to say of fractions with an initial boiling pointof at least 150° C. and a final boiling point reaching to just beforethe initial boiling point of the residue, for example lower than 340°C., or also lower than 370° C.

PRIOR ART

The hydrocracking of heavy petroleum fractions is a very importantrefining process which permits the production, starting from excessheavy charges which are not very exploitable, of lighter fractions suchas gasolines, jet engine fuels and light gas-oils sought by the refinerto adapt his production to the structure of demand. Some hydrocrackingprocesses make it possible to also obtain a highly purified residuewhich can provide excellent bases for oils. In comparison to catalyticcracking, the advantage of catalytic hydrocracking is that it deliversvery high quality middle distillates, jet engine fuels and gas-oils.Conversely, the gasoline produced has a much lower octane index thanthat produced by catalytic cracking.

Hydrocracking is a process which derives its flexibility from three mainelements which are the operating conditions used, the types of catalystsused and the fact that the hydrocracking of hydrocarbon charges can becarried out in one or two-stages. In fact, hydrocracking is a processwhich can assume various forms of which the main ones are:

Single-stage hydrocracking, which firstly and generally comprises anintensive hydrotreatment which has the aim of carrying out an intensivehydrodenitrogenation and an intensive desulphurization of the chargebefore it is sent on to the hydrocracking catalyst proper, in particularin the event it comprises a zeolite. This intensive hydrotreatment ofthe charge involves only a limited conversion of the charge, intolighter fractions, which remains insufficient and must therefore becompleted on the more active hydrocracking catalyst. However, it shouldbe noted that no separation occurs between the two types of catalyst.All of the effluent leaving the reactor is injected onto thehydrocracking catalyst proper and it is only then that a separation ofthe products formed is realized. This single-stage hydrocracking, alsocalled “once through” hydrocracking, has a variant which includes arecycling of the unconverted fraction to the reactor with a view to amore complete conversion of the charge.

Two-stage hydrocracking comprises a first stage which as in the“one-stage” process, has the aim of carrying out the hydrorefining ofthe charge but also achieving a conversion of the latter that isgenerally of the order of 40 to 60%. The effluent from the first stagethen undergoes a separation (distillation) most often calledintermediate separation, which has the aim of separating the conversionproducts from the unconverted fraction. In the second-stage of a 2-stagehydrocracking process, only the fraction of the charge which is notconverted during the first stage is treated. This separation allows atwo-stage hydrocracking process to be more selective in diesel than asingle-stage process. In fact, the intermediate separation of theconversion products prevents their “over-cracking” into naphtha and gasin the second-stage on the hydrocracking catalyst. Furthermore, itshould be noted that the unconverted fraction of the charge treated inthe second-stage generally contains very low NH₃ levels and of organicnitrogen compounds, in general less than 20 ppm by weight or even lessthan 10 ppm by weight. In a standard manner, the 2-stage process can becarried out either with an intermediate separation after hydrorefining,in a process comprising a hydrorefining reactor and a hydrocrackingreactor, or with an intermediate separation between the first and thesecond hydrocracking reactors in a process in which the hydrorefining1^(st) hydrocracking, 2^(nd) hydrocracking reactors are in series.

The hydrocracking catalysts used in the hydrocracking processes are allof the bifunctional type combining an acid function with a hydrogenatingfunction. The acid function is provided by large surfaced supports(generally 150 to 800 m².g⁻¹) presenting a superficial acidity, such ashalogenated (in particular chlorinated or fluorinated) aluminas,combinations of boron and aluminium oxides, amorphous silica-aluminasand zeolites. The hydrogenation function is provided either by one ormore metals of group VIII of the periodic table of elements, or bycombination of at least one metal of group VIB of the periodic table andat least one group VIII metal.

In general, these catalysts are present downstream of the hydrotreatmentreactor in the single-stage hydrocracking processes or in thesecond-stage of the 2-stage hydrocracking processes. However, they canalso be present in the first hydrocracking stage.

The choice of catalysts to be used in the different types ofhydrocracking process and in the different stages in the case of atwo-stage process will depend in particular, on the type of charge to betreated as well as the aim assigned to the hydrocracker: maxi-gasolineor maxi middle distillate (kerosene+gas-oil).

Generally, the balance between the two functions, acid andhydrogenation, is the fundamental parameter which governs the activityand the selectivity of the catalyst. A weak acid function and a stronghydrogenation function give not very active catalysts, working atgenerally high temperatures (greater than or equal to 390° C.), and atlow spatial feed rates (the VVH expressed as a volume of the charge tobe treated per volume unit of the catalyst per hour is generally lowerthan or equal to 2 h-¹) but displaying a very good selectivity asregards middle distillates. Conversely, a strong acid function and aweak hydrogenation function produce active catalysts which do however,display less satisfactory selectivities as regards middle distillates.The search for a suitable catalyst will therefore be centred on ajudicious choice of each of the functions to adjust theactivity/selectivity combination of the catalyst.

Therefore, one of the great advantages of hydrocracking is that itdisplays great flexibility at various levels: flexibility regarding thehydrocracking process to be used, the catalysts used, which lead to aflexibility in the charges to be treated and a diversity in theselectivity of the products obtained.

A first type of conventional catalytic hydrocracking catalysts is basedon low-acidity amorphous supports, such as amorphous silica-aluminas forexample. These systems are more particularly used to produce very highquality middle distillates, and also, when their acidity is very weak,oil bases. These catalysts are in general used in the two-stageprocesses.

The family of the amorphous silica-aluminas is found in low-aciditysupports.

Many hydrocracking catalysts are based on silica-alumina, combinedeither with a group VIII metal or, preferably when the heteroatomictoxin content of the charge to be treated exceeds 0.5% by weight, with acombination of metal sulphides of groups VIB and VIII. These systemshave a very good selectivity in respect of middle distillates, and theproducts formed are of good quality. These catalysts, or the less acidamong them, can also produce lubricating bases. The drawback of all ofthese catalytic systems based on an amorphous support is, as alreadymentioned, their low activity.

Other conventional catalysts comprising the Y zeolite of structural typeFAU, or the beta-type catalysts have a catalytic activity better thanthat of the amorphous silica-aluminas, but displays greaterselectivities in respect of light products. These catalysts are ingeneral used in the “once through” single-stage processes or withrecycling of the unconverted fraction. They have also been used in thesecond-stage of a two-stage hydrocracking process.

According to the prior art, all these 2-stage processes operate in theabsence of ammonia (or its quasi-absence) from the 2^(nd) hydrocrackingreactor, and this is essentially for two reasons. The first reason isthat, in the absence of ammonia, the 2^(nd) hydrocracking reactor canfunction at a lower temperature than the 1^(st)reactor (270–370° C. and300–450° C. respectively). The second reason is that the absence ofammonia allows the use in the 2^(nd) stage of catalysts with noblemetals or metal sulphides. This absence or quasi-absence of ammonia hasalways been proposed and used.

However, U.S. Pat. No. 3,816,296 teaches that it is possible, when usinga catalyst optionally comprising a zeolite, to increase the selectivityin respect of middle distillates of the second-stage of hydrocracking ahydrocarbon charge containing less than 10 ppm by weight organicnitrogen by adding to the latter a quantity of nitrogen (from ammonia oramines with less than 15 carbon atoms) of between 15 and 100 ppm byweight (relative to the charge). The quantity of nitrogen added musttherefore be strictly controlled and maintained within these limits.

Contrary to the prior art, the research work carried out by theapplicant has led to the discovery that, surprisingly, in a 2-stageprocess the quantity of nitrogen added to obtain clearly improvedselectivity in respect of middle distillates could be increased wellabove 110 ppm by weight of nitrogen, whilst retaining good catalyticactivity, even though a catalyst comprising a Y zeolite is used in thesecond-stage of the process.

DETAILED DESCRIPTION OF THE INVENTION

More specifically, the invention describes a 2-stage hydrocrackingprocess of hydrocarbon charges, comprising a first stage including ahydrorefining, an intermediate separation of the converted products, anda second-stage of hydrocracking of at least part of the residue, thesaid second-stage operating in the presence of ammonia in a quantitycorresponding to more than 100 ppm nitrogen and advantageously thequantity of nitrogen is greater than 150 ppm and preferably greater than200 ppm. Generally, it is at the most 1000 ppm, or at the most 800 ppm,or at the most 500 ppm.

The presence of ammonia in these quantities allows significant gains forthe selectivity in respect of middle distillates of the zeoliticcatalyst, a selectivity which therefore becomes comparable to that ofamorphous catalysts containing for example an amorphic silica as acidfunction. The improved selectivity is obtained with reasonable increasesin reaction temperatures whilst preserving the stability of the zeolite,that is to say the duration of the catalyst cycle. It has also beenfound that the selectivity in respect of gas-oil (for example offraction points 250–380° C.) is greater for high quantities of nitrogen(more than 150 ppm, or better still more than 200 ppm by weight).

The additional presence of ammonia is obtained by direct injection ofammonia or also by injection of a nitrogen compound which breaks downinto ammonia in the reaction conditions, the injection taking placedirectly into the reactor and for example at several points of thereactor, or preferably into the charge entering this reactor. Numerousnitrogen compounds can be used, for example aniline.

First Stage

Very varied charges can be treated by the process according to theinvention and generally they contain at least 20% by volume and often atleast 80% by volume of compounds which boil above 340° C.

The charge can be for example LCO (light cycle oil), atmosphericdistillates, distillates under vacuum for example gas-oil from thedirect distillation of the crude or conversion units such as FCC, cokeror visbreaker, as well as charges originating from units extractingaromatics from lubricating oil bases or resulting from the solvent basedremoval of paraffins from lubricating oil bases, or distillatesoriginating from desulphurization or hydroconversion of ATR (atmosphericresidues) and/or RUV (residues under vacuum), or the charge can be adeasphalted oil, or any mixture of the charges previously mentioned. Thelist above is not limitative. In general, the charges have an initialboiling point greater than 340° C., and better still greater than 370°C.

The nitrogen content is usually between 1 and 5000 ppm by weight, moregenerally between 200 and 3000 ppm by weight, and the sulphur contentbetween 0.1 and 5% by weight, more generally between 0.2 and 4%.

In the first stage the charge undergoes at least one hydrorefining(hydrodesulphurization, hydrodenitrogenation, conversion). Standardcatalysts can be used, which contain at least one amorphous support andat least one hydro-dehydrogenating element (generally at least onenon-noble element from the groups VIB and VIII, and most frequently atleast one element from group VIB and at least one non-noble element fromthe VIII group)

In a very advantageous manner, in the two-stage hydrocracking process,according to the invention, the charge to be treated is placed incontact, in the presence of hydrogen, with a hydrorefining catalystcomprising at least one matrix, at least one hydro-dehydrogenatingelement chosen from the group formed by the elements of group VIB andgroup VIII of the periodic table, optionally at least one promoterelement deposited on the catalyst and chosen from the group formed byphosphorus, boron and silicon, optionally at least one element fromgroup VIIA (preferably chlorine and fluorine), and optionally at leastone element from group VIIB (preferably manganese), optionally at leastone element from group VB (preferably niobium).

Preferably, this catalyst contains boron and/or silicon as a promoterelement, plus optionally phosphorus as another promoter element. Theboron, silicon and phosphorus contents are therefore from 0.1–20%,preferably 0.1–15%, and even more advantageously 0.1–10%.

The matrices which can be used on their own or in a mixture are, by wayof non limitative example, alumina, halogenated alumina, silicon,silica-alumina, clays (for example the natural clays such as kaolin orbentonite), magnesium, titanium oxide, boron oxide, zirconia, aluminiumphosphates, titanium phosphates, zirconium phosphates, carbon andaluminates.

The use of matrices containing alumina is preferred, in all these formsknown to a person skilled in the art, and in an even more preferredmanner aluminas, for example gamma alumina.

The role of hydro-dehydrogenating function is preferably fulfilled by atleast one non-noble metal or metal compound from groups VI and VIIIpreferably chosen from among molybdenum, tungsten, nickel and cobalt.Preferably, this function is fulfilled by the combination of at leastone element of GVIII (Ni, Co) with at least one element of group VIB(Mo, W).

This catalyst can advantageously contain phosphorus; in fact it is knownin the prior art that this compound gives the hydrotreatment catalyststwo advantages: an ease of preparation notably during the impregnationof the nickel and molybdenum solutions, and a better hydrogenationactivity.

In a preferred catalyst, the total concentration of oxides of metals ofgroups VI and VIII between 5 and 40% by weight and preferably between 7and 30% and the weight ratio expressed as metal oxide of metal (ormetals) of group VIB to metal (or metals) group VIII is preferablybetween 20 and 1.25 and even more preferably between 10 and 2. Theconcentration of phosphorus oxide P₂O₅ will be lower than 15% by weightand preferably 10% by weight.

Another preferred catalyst which contains boron and/or silicon (andpreferably boron and silicon), generally contains, in % by weightrelative to the total mass of the catalyst, at least one metal chosenfrom the following groups and with the following contents:

-   -   3 to 60%, preferably from 3 to 45% and in an even more preferred        manner from 3 to 30% of at least one metal of group VIB and        optionally,    -   0 to 30%, preferably from 0 to 25% and in an even more preferred        manner from 0 to 20% of at least one metal of group VIII,    -   the catalyst also containing at least one support chosen from        the following groups with the following contents:        -   0 to 99%, advantageously 0.1 to 99%, preferably from 10 to            98% and in an even more preferred manner from 15 to 95% of            at least one amorphous or poorly crystallized matrix,

the said catalyst being characterised in that it also contains,

-   -   0.1 to 20%, preferably from 0.1 to 15% and in an even more        preferred manner from 0.1 to 10% boron and/or 0.1 to 20%,        preferably from 0.1 to 15% and in an even more preferred manner        from 0.1 to 10% silicon and optionally,    -   0 to 20%, preferably from 0.1 to 15% and in an even more        preferred manner from 0.1 to 1 0% phosphorus, and optionally        also,    -   0 to 20%, preferably from 0.1 to 15% and in an even more        preferred manner from 0.1 to 10% of at least one element chosen        from group VIIA, preferably fluorine.

In a general manner, formulas having the following atomic ratios arepreferred:

-   -   an atomic ratio of the group VIII metal to the group VIB metals        of between 0 and 1,    -   an atomic ratio B/metals of group VIB comprised between 0.01 and        3,    -   an atomic ratio Si/metals of group VIB comprised between 0.01        and 1.5,    -   an atomic ratio P/metals of group VIB comprised between 0.01 and        1,    -   an atomic element ratio of the VIIA/metals group of group VIB        comprised between 0.0 1 and 2.

In terms of the hydrogenation of the aromatic hydrocarbons and ofhydrodenitrogenation and hydrosulphurization, such a catalyst has agreater activity than the catalytic formulas without boron and/orsilicon, and also has an activity and selectivity as regardshydrocracking greater than the catalytic formulas known in the priorart. The catalyst with boron and silicon is particularly advantageous.Without wishing to be tied down by any particular theory, it seems thatthis particularly high activity of the catalysts with boron and siliconis due to the strengthening of the acidity of the catalyst by the jointpresence of boron and silicon on the matrix which induces on one hand animprovement in the hydrogenating, hydrodesulphurizing, hydrodeazotingproperties and on the other hand an improvement in the hydrocrackingactivity in comparison with catalysts normally used in the hydrorefiningreactions of hydroconversion.

The preferred catalysts are the NiMo and/or NiW catalysts on alumina,and also the NiMo and/or NiW catalysts on alumina doped with at leastone element included in the group of atoms formed by phosphorus, boron,silicon and fluorine, or the NiMo and/or NiW catalysts on silica-amina,or on silica-alumina-titanium oxide doped or not by at least one elementincluded in the group of atoms formed by phosphorus, boron, fluorine andsilicon.

Another particularly advantageous type of catalyst (particularly asregards improved activity) in hydrorefining contains a partiallyamorphous Y zeolite, this catalyst will be described later in thesecond-stage.

In a general manner, the 1^(st)-stage hydrorefining catalyst contains:

-   -   5–40% by weight of at least one non-noble element of the groups        VIB and VIII (% oxide)    -   0–20% of at least one promoter element chosen from among        phosphorus, boron, silicon (%oxide), preferably 0.1–20%;        advantageously boron and/or silicon are present, and optionally        phosphorus.    -   0–20% of at least one element from group VIIB (manganese for        example)    -   0–20% of at least one element from group VIIA (fluorine,        chlorine for example)    -   0–60% of at least one element from group VB (niobium for        example)    -   0.1–95% of at least one matrix, and preferably alumina

The catalysts described above are generally used to provide thehydrorefining also called the hydrotreatment stage. This hydrorefiningstage can be followed by an intermediate separation (the unconvertedeffluent then goes into the second-stage), or all of the effluentleaving the hydrorefining stage is treated by a 1^(st) stagehydrocracking catalyst.

This first stage hydrocracking is carried out for example in anotherreactor or in an additional catalyst bed in the reactor where thehydrorefining takes place. A pre-cracking is then achieved which makesit possible to reach the desired rates of conversion in the first stage.In this first stage hydrocracking stage, the catalyst used possesses atleast one hydrodehydrogenating function, a greater acidity making itpossible to supplement the pre-cracking. This greater acidity can beprovided by an acid solid such as an amorphous silica-alumina or azeolite. The hydrocracking catalyst is identical to that of thesecond-stage (as described later) or different and is preferablyzeolitic.

Prior to injection of the charge, the catalysts used in the processaccording to the present invention are preferably subjected beforehandto a sulphurization treatment making it possible to transform, at leastin part, the metallic types to sulphur before they are brought intocontact with the charge to be treated. This activation treatment bysulphurization is well known to a man skilled in the art and can becarried out by any method already described in literature either insitu, that is to say in the reactor, or ex situ.

A standard method of sulphurization well known to a person skilled inthe art consists of heating in the presence of hydrogen sulphide (pureor for example under flux of a hydrogen hydrogen sulphide mixture) to atemperature of between 150 and 800° C., preferably between 250 and 600°C., generally in a crossed bed reaction zone.

In the first stage of the process, the charge is brought into contact,in the presence of hydrogen, with at least one catalyst as previouslydescribed, at a temperature of between 330 and 450° C., preferably360–420° C., preferably 360–420° C., under a pressure of between 5 and25 MPa, preferably lower than 20 MPa, the spatial velocity being between0.1 and 6 h⁻¹, preferably 0.2–3h⁻¹, and the quantity of hydrogenintroduced is such that the per litre of hydrogen/litre of hydrocarbonvolume ratio is between 100 and 2000 I/I.

During this stage, a substantial reduction in the level of organicnitrogenous and sulphurous compounds and condensed polycyclic aromatichydrocarbons is achieved. In these conditions, the majority of theorganic nitrogenous and sulphurous products of the charge are alsotransformed into H₂S and into NH₃. This operation therefore makes itpossible to eliminate two types of compounds which are known to beinhibitors of the zeolitic catalyst.

In the process according to the invention the level of nitrogen organiccompounds in the effluent leaving the first stage is lower than 20 ppmby weight and preferably lower than 10 ppm by weight.

This first stage (including hydrorefining and optionally hydrocracking)also makes it possible to carry out a pre-cracking of the charge to betreated.

Advantageously, this adjustment can be carried out by utilizing thenature and the quality of the catalyst(s) used in the first stage and/orthe operating conditions of this first stage. In the process accordingto the invention the conversion, during the first stage, in terms ofproducts with boiling points lower than 340° C., and better, lower than370° C., is greater than 20% and preferably greater than 30% and in aneven more preferred manner between 40 and 60%.

Intermediate Separation

The effluent resulting from this first stage is sent to a means ofseparation (separating flask for example) the purpose of which is tocarry out a separation of the ammonia (NH₃) and the hydrogen sulphide(H₂S) produced during this first stage. The hydrocarbon effluentproduced by this separation will undergo an atmospheric distillation,and in some cases the combination of an atmospheric distillation and adistillation under vacuum. The purpose of the distillation is carryingout a separation between the converted hydrocarbon products, that is tosay generally having boiling points lower than 340° C. (and better,lower than 370° C.) and an unconverted liquid fraction (residue).

Advantageously, distillation can be carried out at atmospheric pressureto obtain several converted fractions (petrol, kerosene, gas-oil forexample, with a boiling point of at the most 340° C.) and a residuefraction (for example with an initial boiling point greater than 340° C.or even greater than 370° C.).

To improve the separation, distillation under vacuum can be added. Thiswill be the case for example in order to distil diesel more effectively,or also when it is wished to remove a heavy fraction of the residue fromthe 2^(nd-)stage passage. The liquid fraction, residue, containingproducts the boiling point of which is greater than 340° C. or indeedeven greater than 370° C. and resulting from the distillation is atleast in part and preferably totally introduced into the second-stage ofthe process according to the invention.

Second-Stage

The residue fraction resulting from the intermediate separation and sentto the second-stage is said to be “clean” that is to say it containsless than 10 ppm by weight of organic nitrogen and less than 10 ppm byweight of organic sulphur, that is to say nitrogen and sulphur includedin organic compounds.

According to the invention, at least one nitrogen compound decomposablein ammonia, in the conditions of the second-stage or ammonia direct isadded to the charge or injected into the second-stage reactor.

Among the usable nitrogen compounds, there may be cited by way ofexample and in a non-exhaustive manner, aromatic amines (aniline forexample), aliphatic amines (nButylamine for example), pyrroles,pyridines; ureas; nitrated, nitrous or nitroso derivatives; primary,secondary or tertiary amines; compounds with ammonium.

The quality of ammonia (NH₃) added to the reactor(s) of the second-stageof the hydrocracking process according to the invention is such that inthe said reactor(s) the nitrogen content by weight, expressed in ppm byweight (parts per million) relative to the charge entering the saidreactor(s) is greater than 110 ppm and preferably greater than 150 ppmand in an even more preferred manner greater than 200 ppm by weight.

Furthermore, the second-stage catalyst being, for the reaction,sulphurous, it is advisable to keep it in contact with a partial H₂Spressure sufficient to avoid its desulphurization in the presence ofhydrogen and at the reaction temperatures. To this end, and in astandard manner, hydrogen sulphide or at least one sulphur compoundwhich decomposes in H₂S in the conditions of the second-stage is addedto the charge or directly into the reactor.

There may be cited, as a sulphur compound, dimethyldisulphide (DMDS),carbon disulphide (CS₂), organic polysulphides, mercaptans, sulphides,disulphides, oxygenated sulphur compounds, elemental sulphur, dissolvedand/or partially in suspension.

The quantity of hydrogen sulphide (H₂S) added into the reactors of thesecond-stage of the hydrocracking process according to the inventioncorresponds to a sulphur content by weight, expressed in ppm (parts permillion) relative to the charge entering the said reactor(s) greaterthan 20 ppm and preferably greater than 50 ppm and in an even morepreferred manner greater than 200 ppm.

The NH₃ and H₂S quantities can be regulated throughout the reaction bythe operator. When the second-stage comprises several reactors, theaddition takes place in at least one reactor (into the charge ordirectly into the reactor).

Advantageously., the ammonia (NH₃) and the hydrogen sulphide (H₂S)injected into the second-stage of the hydrocracking process according tothe invention come from the recycling of at least part of the ammoniaand hydrogen sulphide produced in the first stage of the process andobtained during the intermediate separation.

Advantageously, at least part of the ammonia produced in the first stageand separated will be used as a source of ammonia. This can be the gascontaining NH₃, H₂S, H₂ and the light hydrocarbons separated for examplein a gas-liquid separator. Advantageously, it can be a more purified gasobtained after washing the first-stage effluent with water, separationof the aqueous phase and stripping of this aqueous phase in such a wayas to produce an ammonia gas containing a little hydrogen and a fewlight gases. This latter embodiment will be described later, withreference to the figures.

The operating conditions used in the second-stage of the processaccording to the invention are: a temperature greater than 200° C.,often between 250–480° C., advantageously between 320 and 450° C.,preferably between 330 and 425° C., at a pressure greater than 0.1 MPa,often between 5 and 25 MPa, preferably lower than 20 MPa and even moreadvantageously greater than 9 MPa, or better than 10 MPa, the spatialvelocity being between 0.1 and 20h⁻¹ and preferably 0.1–6h⁻¹, preferably0.2–3h⁻¹, and the quantity, of hydrogen introduced is such that thelitre of hydrogen/litre of hydrocarbon volume ratio is between 80 and5000 l/l and most frequently between 100 and 2000 l/l.

These operating conditions used in the second-stage of the processaccording to the invention make it possible to obtain conversion ratesper passage, in terms of products having boiling points lower than 340°C. and better, lower than 370° C., greater than 30% and in an even morepreferred manner between 40 and 60%.

The second-stage catalyst comprises at least one Y zeolite, at least onematrix and a hydro-dehydrogenating function. Optionally, it can alsocontain at least one element chosen from among boron, phosphorus andsilicon, at least one element from G VIIA (chlorine, fluorine forexample), at least one element from G VIIB (manganese for example), atleast one element from G VB (niobum for example).

The catalyst contains at least one porous or poorly crystallizedoxide-type mineral matrix. There may be cited, as a non-limitingexample, aluminas, silicas, silica-aluminas, aluminates, boronaluminium-oxide, magnesium, silica-magnesium, zirconium, titanium oxide,clay, on their own or in a mixture.

The hydro-dehydrogenating function is generally fulfilled by at leastone element from group VI B (for example molybdenum and/or tungsten)and/or at least one element from the non-noble VIII group (for examplecobalt and or nickel) of the periodic table of elements.

A preferred catalyst essentially contains at least one metal of groupVI, and/or at least one non-noble metal from group VIII, the Y zeoliteand alumina.

An even more preferred catalyst essentially contains nickel, molybdenum,a Y zeolite and alumina.

In a preferred manner, the catalyst contains at least one element chosenfrom the group formed by boron, silicon and phosphorus. Advantageously,the catalyst optionally contains at least one element from group VIIA,preferably chlorine and fluorine, optionally at least one element fromgroup VIIB (manganese for example), optionally at least one element fromgroup VB (niobium for example).

The boron, silicon and/or phosphorus can be in the matrix, the zeoliteor are preferably deposited on the catalyst and are therefore mainlylocated on the matrix. A preferred catalyst contains B and/or Si asdeposited promoter element preferably also with phosphorus promoter. Thequantities introduced are from 0.1–20% by weight of catalyst calculatedas oxide.

The element introduced, and in particular the silicon, mainly located onthe matrix of the support can be characterised by techniques such as theCastaing microprobe (distribution profile of the various elements),electron microscopy by transmission coupled with an X analysis of thecomponents of the catalysts, or by the establishment of a distributioncartography of the elements present in the catalyst by electronmicroprobe.

Generally, the 2^(nd)-stage catalyst advantageously contains:

-   -   0.1–80% by weight zeolite Y    -   0.1–40% by weight of at least one element from groups VIB and        VIII (% oxide)    -   01–99.8% by weight of matrix (% oxide)    -   0–20% by weight of at least one element chosen from the group        formed by P, B, Si (% oxide), preferably 0. 1–20%    -   0–20% by weight of at least one element from group VIIA,        preferably 0.1–20%    -   0–20% by weight of at least one element from group VIIB,        preferably 0. 1–20%    -   0–60% by weight of at least one element from group VB,        preferably 0.1–60%

As far as the silicon is concerned, in the 0–20% range only the addedsilicon, and not that of the zeolite, is counted.

The zeolite can optionally be doped by metallic elements such as forexample metals from the family of rare earths, in particular lanthanumand cerium, or noble or non-noble metals from group VIII, such asplatinum, palladium, ruthenium, rhodium, iridium, iron and other metalssuch as manganese, zinc and magnesium.

Different Y zeolites can be used.

A particularly advantageous H—Y acid zeolite is characterised bydifferent specifications: a global SiO₂/Al₂O₃ molar ratio of betweenabout 6 and 70 and in a preferred manner between about 12 and 50: asodium content lower than 0.15% weight determined on zeolite calcined at1 100° C.; a crystalline parameter with a lattice-cell between24.58×10⁻¹⁰ m and 24.24×10⁻¹⁰ m and in a preferred manner between24.38×10⁻¹⁰ m and 24.26×10⁻¹⁰ m; a sodium ion uptake capacity CNa,expressed in grams of Na per 100 grams of modified zeolite, neutralisedthen calcined, greater than about 0.85; a specific surface determined bythe B.E.T. method greater than about 400 m²/g and preferably greaterthan 550 m²/g, a water vapour adsorption capacity at 25° C. for apartial pressure of 2.6 torrs (i.e. 34.6 MPa), greater than about 6%,and advantageously, the zeolite has a pore distribution, determined byphysisorption of nitrogen, of between 5 and 45% and preferably between 5and 40% of the total porous volume of the zeolite contained in poreswith a diameter of between 20×10⁻¹⁰ m and 80×10⁻¹⁰ m and between 5 and45% and preferably between 5 and 40% of the total porous volume of thezeolite contained in pores with a diameter greater than 80×10⁻¹⁰ m andgenerally lower than 1000×10⁻¹⁰ m, the rest of the porous volume beingcontained in pores with a diameter lower than 20×10⁻¹⁰ m.

A preferred catalyst using this type of zeolite contains a matrix, atleast one dealuminized Y zeolite possessing a crystalline parameter ofbetween 2.424 nm and 2.455 nm preferably between 2.426 and 2.438 nm, aglobal SiO₂/Al2O₃ molar ratio greater than 8, an alkaline-earth metalsor alkali cations and/or rare earths cations content such as the(n×M^(n+))/Al atomic ratio is lower than 0.8 preferably lower than 0.5or even 0.1, a specific surface determined by the B.E.T method greaterthan 400 m²/g, preferably greater than 550m²/g, and a water absorptioncapacity at 25° C. for a P/Po value of 0.2, greater than 6% by weight,the said catalyst also comprising at least one hydro-dehydrogenatingmetal, and silicon deposited on the catalyst.

In an advantageous embodiment according to the invention, there is usedfor the hydrocracking (second-stage and/or optionally first-stage) acatalyst comprising a partially amorphous Y zeolite.

By partially amorphous Y zeolite is meant a solid presenting:

-   -   i/ a rate of peaks which is lower than 0.40 preferably lower        than about 0.30    -   ii/a crystalline fraction expressed relative to a reference Y        zeolite in sodium form (Na) which is lower than about 60%,        preferably lower than about 50%, and determined by X-ray        diffraction.

Preferably, the solid, partially amorphous Y zeolites involved in thecomposition of the catalyst according to the invention have at least one(and preferably all) of the following characteristics:

-   -   iii/ a global Si/AI ratio greater than 15, preferably greater        than 20 and lower than 150,    -   iv/ an Si/AI^(iv) framework ratio greater than or equal to the        global Si/AI ratio,    -   v/a porous volume at least equal to 0.20 ml/g, of solid        material, a fraction of which, between 8% and 50%, is        constituted by pores with a diameter of at least 5 nm        (nanometre) i.e. 50 Å.    -   vi/ a specific surface of 210–800 m²/g, preferably 250–750 m²/g        and advantageously 300–600 mn²/g

The rate of peaks and the crystalline fractions are determined by X-raydiffraction, using, a procedure derived from ASTM methodD3906-97<<Determination, of Relative X-ray Diffraction Intensities ofFaujasite-Type-Containing Materials>>. Reference may be made to thismethod for the general conditions of use of the process and, inparticular, for the preparation of samples and references.

A diffractogram is composed of lines characteristic of the crystallisedfraction of the sample and of a trough, caused essentially by thediffusion of the amorphous or microcrystalline fraction of the sample (aweak diffusion signal is linked to the apparatus, air, sample holderetc.) The rate of peaks of a zeolite is the ratio, in a predefinedangular zone (typically 8 to 40° 2θ when the. Kα radiation of copper,1=0.154 nm is used), the area of the, zeolite rays (peaks) for theglobal area of the diffractogram (peaks+trough). Thispeaks/peaks+trough) ratio is proportional to the quantity ofcrystallized zeolite in the material. To estimate the crystallinefraction of a sample of Y zeolite, the rate of peaks of the sample willbe compared with that of a reference considered to be 100% crystallized(NaY for example). The rate of peaks of a perfectly crystallized NaYzeolite is of the order of 0.55 to 0.60.

The rate of peaks of a standard USY zeolite is from 0.45 to 0.55, itscrystalline fraction relative to a perfectly crystallized NaY is from 80to 95%. The rate of peaks of the solid forming the subject of thepresent invention is lower than 0.4 and preferably lower than 0.35. Itscrystalline fraction is therefore lower than 70%, preferably lower than60%.

The partially amorphous zeolites are prepared according to thetechniques generally used for dealuminization, from commerciallyavailable Y zeolites, that is to say those which generally have highcrystallinity levels (at least 80%). More generally zeolites can be usedwhich have a crystalline fraction of at least 60%, or at least 70%.

The Y zeolites generally used in hydrocracking catalysts aremanufactured by modifying commercially available Na-Y zeolites. Thismodification allows zeolites to be obtained which are called stabilized,ultra-stabilized or even dealuminized. This modification is carried outby at least one of the dealuminzation techniques, and for example byhydrothermal treatment, acid attack. Preferably, this modification iscarried out by a combination of three types of operations known to aperson skilled in the art: hydrothermal treatment, ion exchange and acidattack.

Another particularly useful zeolite is a globally non-dealuminized andvery acid zeolite.

By non-globally dealuminized zeolite is meant a Y zeolite (FAU,structural type faujasite) according to the detailed nomenclature in“Atlas of zeolites structure types”, W. M. Meier, D. H. Olson and Ch.Baerlocher, 4th revised Edition 1996, Elsevier. The crystallineparameter of this zeolite may have diminished through extraction of thealuminas of the structure or framework during the preparation but theglobal SiO₂/Al2O₃ ratio has not changed as the aluminas were notextracted chemically. Such a non-globally dealuminized zeolite thereforehas a silicon and alumina composition expressed by the global SiO₂/Al₂O₃ratio equivalent to the starting non-globally dealuminized Y zeolite.The values of the parameters (SiO₂/Al₂O₃ ratio and crystallineparameter) are given later on. This non-globally dealuminized Y zeolitecan be in the hydrogen form or be at least partially exchanged withmetallic cations, for example with the aid of cations of thealkaline-earth metals and/or cations of rare earth metals with atomicnumbers from 57 to 71 inclusive. A zeolite without rare earths andalkaline-earths will be preferred, likewise for the catalyst.

The non-globally dealuminized Y zeolite generally has a crystallineparameter greater than 2.43 8 nm, a global SiO₂/Al₂O₃ ratio lower than8, a framework SiO₂/Al₂O₃ molar ratio lower than 21 and greater than theglobal SiO₂/Al₂O₃ ratio.

The non-globally dealuminized zeolite can be obtained by any treatmentwhich does not extract the aluminas from the sample, such as for examplethe treatment with water vapour, treatment with SiCI₄.

Another type of catalyst which is advantageous for hydrocrackingcontains an acid amorphous oxide matrix of alumina type doped withphosphorus, a non-globally dealuminized and very acid Y zeolite andoptionally at least one element of group VIIA and in particularfluorine.

The invention is not limited to the cited and preferred Y zeolites, butother types of Y zeolites can be used in this process.

Prior to the injection of the charge into the second-stage of theprocess according to the present invention, the catalyst is subjected toa sulphurization treatment making it possible to transform, at least inpart, the metallic types into sulphur before they are brought intocontact with the charge to be treated. This treatment of activation bysulphurization is well known to a person skilled in the art and can becarried out by any method already described in literature eitherin-situ, that is to say in the reactor, or ex-situ.

A standard method of sulphurization well known to a person skilled inthe art consists of heating in the presence of hydrogen sulphide (pureor for example under flux of a hydrogen/hydrogen sulphide mixture) to atemperature between 150 and 800° C., preferably between 250 and 600° C.,generally in a crossed bed reaction zone.

The effluent leaving the second-stage of the hydrocracking processaccording to the invention, is subjected to a so-called final separation(for example by atmospheric distillation optionally followed by adistillation under vacuum), in order to separate the gases (such asammonia (NH₃) and hydrogen sulphide (H₂S), as well as the other lightgases present, the hydrogen and optionally the conversion products . . .). At least one liquid residue fraction is obtained containing productsthe boiling point of which is generally greater than 340° C., which isat least in part recycled in the second-stage of the process.

Advantageously (as shown in the figures), the final separation iscarried out with the means of intermediate separation when thesecomprise an atmospheric distillation and optionally distillation undervacuum.

The invention thus also relates to an installation to carry out atwo-stage hydrocracking process, the installation comprising:

-   -   at least one first stage hydrorefining reactor (2) comprising at        least one catalyst bed to carry out hydrorefining of the charge,    -   at least one duct (1) to introduce the charge into the first        reactor of the first stage, which is a hydrorefining reactor, at        least one duct (3) to carry the hydrogen to said reactor and at        least one exit duct (4) for the effluent of the last reactor of        the first stage.    -   at least one gas-liquid separator (5) to separate the effluent        leaving the first stage, at least one gas leaving per duct (6),    -   at least one column (8) to separate the products converted        during the first stage and thus obtain a residue.    -   at least one second-stage hydrocracking reactor (14) comprising        at least one catalyst bed to carry out the hydrocracking of at        least part of the said residue,    -   at least one duct (16) for the introduction of hydrogen into at        least the first hydrocracking reactor of the second-stage, at        least one exit duct (17) for the second-stage effluent from the        last second-stage reactor,    -   at least one means of separation (18) to separate the gases of        the effluent leaving the last reactor of the second-stage, and        at least one column to separate-at least part of said effluent,        the converted products and a residue,    -   at least one duct (13) to recycle at least part of the residue        into the 2^(nd) stage hydrocracking reactor (14),    -   the installation also comprises,    -   at least one duct (16) for the introduction of ammonia at least        into the first reactor of the second-stage.

DESCRIPTION OF THE FIGURES

The invention will be illustrated in the figures:

FIG. 1 represents a simplified diagram of the process and theinstallation

FIG. 2 represents a preferred embodiment

FIGS. 3A, 3B, 3C represent various possibilities for the introduction ofammonia or a precursor of ammonia

In FIG. 1, the charge to be treated enters via pipe (1) into the firststage hydrorefining reactor (2) containing at least one hydrorefiningcatalyst bed. It is mixed with hydrogen carried by a pipe (3). This canbe a make-up hydrogen and/or recycled hydrogen, as described in FIG. 2.

The gases are separated from the effluent leaving the first stage viathe pipe (4) into a gas-liquid separator (5), a separator flask forexample. The gases are recovered by a pipe (6) and the resulting liquideffluent by a pipe (7).

According to the 2-stage process, the liquid effluent is then subjectedto an intermediate separation for example in a column (8) so as toseparate the converted products which leave in FIG. 1 via the pipes (9)for the light hydrocarbons (C1–C4), (10) for petrol, (11) for kerosene,(12) for gas-oil.

The unconverted effluent (residue) which leaves from the bottom of thecolumn (8) by the pipe (13) is sent at least in part into thesecond-stage reactor (14) containing at least one bed of hydrocrackingcatalyst.

Ammonia or an ammonia precursor compound is added to the incomingresidue by a pipe (15), and hydrogen (make-up and/or recycled) by a pipe(16).

The gases are separated (pipe 20) into a gas-liquid separator (18) fromthe effluent leaving the second-stage by a pipe (17). The resultingliquid, leaving by a pipe (19), is generally at least partly recycledinto the 2-stage process and preferably into the column (8) so as toseparate the products converted during the second-stage. Another part ofthe liquid cannot be recycled and is removed from the recycling loop,which is called the “bleed” or the purge.

Advantageously, the separators (5) and (18) are supplied with water, andthe 3 phases; gaseous, aqueous and organic, are then separated. Thegaseous phase essentially comprises hydrogen and constitutes therecycling hydrogen which can very advantageously be used to carryhydrogen into the first and second-stage reactors. In the aqueous phase,the ammonium sulphide is dissolved, and in this manner most of the NH₃and H₂S is eliminated from the recycling gas. The organic phaseessentially contains the hydrocarbon products and is sent into thecolumn (8).

FIG. 2 will show these separators in more detail.

The charge which enters via the pipe (1) (refer to FIG. 2 for thedescription which follows), is sent for example into a first stage feedflask (22), to be taken up there by the 1^(st) stage feed pump (23). Itis mixed with the make-up hydrogen carried by the pipe (24) andoptionally with the 1^(st) stage recycling gas introduced by the duct(25) compressed by the make-up compressor (26) and the recyclingcompressor (27) respectively. The mixture is advantageously sentsuccessively into a series of 1^(st)-stage heat exchangers (28), theninto the 1^(st)-stage oven (29) to be brought to the reactiontemperature.

It is then introduced by a pipe (30) into one or more 1^(st) stagereactors (31) where a hydrorefining takes place optionally followed by ahydrocracking.

The reactor comprises one or more fixed catalytic beds, optionallyseparated by quench injections (cooling fluid, generally hydrogen). Theeffluent leaving the reactor via the pipe (32), containing in particularthe molecules of ammonia NH₃ and hydrogen sulphide H₂S produced, ismixed with washing water introduced by the duct (33). The mixture iscooled in the series of heat exchangers (28) followed optionally by acooling tower, in order to be collected in a gas-liquid separator flask(34).

3 phases are recovered from this flask:

-   -   a vapour phase, which can be partially purged via a duct (36),        and of which at least part can be sent into the reactor by means        of the recycling compressor (27) and the duct (25), another part        being able to be sent into the second-stage according to the        process by the duct (44),    -   a hydrocarbon liquid phase (containing the product of the first        stage of hydrocracking) which leaves via duct (37),    -   and aqueous phase leaving via duct (38), and containing        dissolved the ammonium sulphide produced by the reaction:        NH₃+H₂S→NH₄HS

The hydrocarbon liquid 37 is introduced into a distillation line 39.This line consists of one or more distillation columns, and makes itpossible to recover the gases, petrol, kerosene and diesel via the pipes40 a, 40 b, 40 c, and 40 d respectively. As to the product unconvertedby the reaction (residue), it is recovered at the bottom of the column(39), and sent via the pipe (41) into the 2^(nd) stage feed flask (42)to be taken up by the second-stage feed pump (43). After mixing with thesecond-stage recycled hydrogen carried by the pipe (44) through therecycling compressor (27), this fluid is heated by a group of exchangers(45), then an oven (46), to finally be introduced by the pipe (47) intothe second-stage reactor (48). Make-up hydrogen can also, if needed, beintroduced.

The effluent from this second-stage reactor leaving by pipe (49) is atleast partially cooled in the series of exchangers (45), and is sentinto the separator flask (34) common to the two-stages.

Middle distillates (kerosene, petrol, gas-oil) and optionally a heavierfraction recovered by a pipe (54) (bleed) on the exit pipe (41) from thefinal separation unit (column (39) common to the intermediateseparation) are thus obtained as exploitable hydrocarbon products.

In the described diagram, the make-up compressor (26), the recyclingcompressor (27) and the separator flask (34) are common to thetwo-stages. A particular heat exchange system has been described by wayof example in FIG. 2, but all other arrangements are suitable.

According to the process considered, secondary details can vary, such asthe relative injection position of the charge, recycling gas andhydrogen make-up gas, the number and the arrangement of the heatexchangers, or the number of reactors, compressors or separating flasks.The two hydrocracking stages can have a common, or separate, recyclingcompressor and separator flask. These details do not have any effect onthe invention described here.

Furthermore, FIG. 2 shows a single 1^(st) stage reactor which istherefore a hydrorefining reactor, but several reactors can be used,which can include one or more hydrocracking reactors.

The addition of ammonia according to the invention can be carried outaccording to various methods.

In a method illustrated in FIG. 2, the necessary quantity of ammonia isinjected in the form of a liquid containing a nitrogen compound. Thiscompound is chosen in such a way that in the temperature and pressureconditions inside the reactor, and in the presence of hydrogen, itundergoes a decomposition into ammonia NH₃. A compound completelydecomposing to NH₃ and hydrocarbons will be preferred. Among the usablecompounds, aniline or any other compound having the same function in thereaction may be cited.

The decomposition reaction of aniline is written as:

This compound is introduced according to FIG. 2 via a duct (50) into thefeed flask (42) of the unit. The injection pump (51) and the flask (52)containing nitrogenous liquid fed via duct (53) with nitrogenous liquidcompound have also been represented by way of illustration.

The nitrogenous liquid compound can also be introduced at any point ofthe unit located upstream of the reactor (48), and for example, betweenthe pump (43) and the introduction of hydrogen via the duct (44).

In another method, it is in gaseous form that the nitrogenous compoundis introduced into the reaction section. In order to do this, a gascontaining ammonia must be available. It is very advisable that theconcentration of ammonia in this gas be as high as possible, preferablymore than 5% by volume.

In FIGS. 3A, 3B, 3C different methods of introduction have beenrepresented, the other elements of the figures not included belowcorresponding to those in FIG. 2.

According to FIGS. 3A, 3B and 3C, this gas is injected into the reactionsection via a duct (57), generally by means of a compressor (56), thegas being carried to the compressor via a duct (58), a gas flask (55)being able to be provided. The point of injection can be placed at theintake of the recycling compressor (FIG. 3A), at any point in thehigh-pressure section (FIG. 3B), and/or in the feed flask (FIG. 3C). Thelast method is preferred, as it makes it possible to minimise the costof the ammonia compressor.

More generally, it can be said that the duct (57) introducing theammonia discharges into the duct (44) recycling the gas coming from thegas-liquid separator in the 2^(nd) stage hydrocracking reactor.

Or, according to figure BC, the duct (57) introducing the ammoniadischarges into the duct (41) carrying the residue into the 2^(nd) stagereactor.

The following examples illustrate the present invention without howeverlimiting its scope.

EXAMPLE 1 Preparation of a 2^(nd) Stage Hydrocracking CatalystContaining a Y Zeolite

A dealuminized USY zeolite with a global Si/AI molar ratio equal to15.2, an Si/AI framework ratio of 29, a crystalline parameter at 24.29 Åcontaining 0.03% by weight of Na, with a crystalline fraction of 85% isused in this example to prepare the hydrocracking catalyst. The supportof the hydrocracking catalyst containing this Y zeolite is manufacturedin the following manner:

20 grams of the Y zeolite described above are mixed with 80 grams of amatrix composed of ultra-fine tabular boehmite or alumina gel marketedunder the name SB3 by the company Condéa Chemie GmbH. This powdermixture was then mixed with an aqueous solution containing 66% nitricacid by weight then kneaded for 15 minutes. At the end of this kneading,the paste obtained is passed through a die having cylindrical orificeswith a diameter equal to 1.4 mm. The extruded material is then driedovernight at 120° C. under air then calcined at 550° C. under air. Theextruded material support, containing the Y zeolite, is impregnated drywith an aqueous solution of a mixture of ammonium heptamolybdate, nickelnitrate and orthophosphoric acid, dried overnight at 120° C. under airand finally calcined under air at 550° C. The oxide content by weight ofthe NiMoPY catalyst that were obtained are shown in table 1.

TABLE 1 Characteristics of the catalyst Catalyst NiMoPY MoO₃ (% weight)14.1 NiO (% weight) 3.0 P₂O₅ (% weight) 4.5 SiO₂ (% weight) 14.3 Make upto 100% AC₂O₃ (% weight) 64.1

EXAMPLE 2 Preparation of the Second-Stage Charge

The charge of the second-stage is produced by hydrotreatment of adistillate under vacuum on an HR360 catalyst marketed by Procatalyse inthe presence of hydrogen, at a temperature of 395° C. and at the hourlyspatial velocity of 0.55h-1. The conversion into products at 380° C. isabout 50% by weight. After a separation stage, the 380°C.+ fraction iscollected and will serve as a charge for the second-stage

The physico-chemical characteristics of this charge are the following:

TABLE 2 characteristics of the second-stage charge Density (20/4) 0.853Sulphur (ppm by 2.5 weight) Nitrogen (ppm by 1.4 weight) Simulateddistillation Initial point 322° C.  5% point 364° C. 10% point 383° C.50% point 448° C. 90% point 525° C. Final point 589° C.

EXAMPLE 3 Test in the Presence of NH₃ According to the Invention

The charge prepared in example 2 is injected into the 2^(nd) stagehydrocracking test unit which comprises a fixed-bed reactor, withascending circulation (<<up-flow>>) of the charge, into which 50 ml ofcatalyst prepared in example 1 is introduced. Before the injection ofthe charge the catalyst is sulphurized with a gas-oil+dimethyldisulphide(DMDS)+aniline mixture to 350° C. Once the sulphurization has beencarried out, the charge described in table 2 can be treated. Theoperating conditions of the test unit are the following:

TABLE 3 Operating conditions Total pressure 14 Mpa Catalyst 50 mlTemperature 320–420° C. Hydrogen flow 50 I/h rate Charge flow 50 ml/hrate

There are added to the charge described in table 2, a quantity ofaniline corresponding to a nitrogen content of 500 ppm by weight and aquantity of DMDS corresponding to a sulphur content of 2000 ppm byweight. The aniline injected into the reactor in the presence of thecatalyst, and in the catalytic operating conditions described in table3, will decompose leading to the formation of ammonia NH, and the DMDSto that of H₂S.

The catalytic performances obtained in these conditions are described intable 4 of this example. The catalytic performances are expressed by thetemperature needed to reach a crude conversion rate of 70% and by thecrude selectivity in respect of 150–380° C. middle distillates for thisconversion these catalytic performances are measured on the catalystonly after a stabilisation period, generally at least 48 hours, had beenobserved.

The crude conversion CC is taken to be equal to:CC=% by weight at 380° C.⁻ of the effluent

The crude selectivity CS of middle distillates is taken to be equal to:CS=[weight of the fraction (150° C.–380° C.) of effluent]/[weight of the380° C.-fraction of the effluent]*100 in % by weight

EXAMPLE 4 Comparative Test

This test is carried out in the same conditions as that of example 3,except for the quantity of aniline added which corresponds to 100 ppm byweight nitrogen.

TABLE 4 Results Nitrogen content of the CS of 150/380 middle charge (ppmby Reaction temperature to distillates at 70% weight) reach 70% of CC ofCC 100 356 61 500 378 66

Table 4 shows that the use of a catalyst comprising a Y zeolite, in theconditions of the two-stage hydrocracking process according to theinvention, leads to an iso-conversion of 70% by weight, with aselectivity in respect of middle distillates (150–380° C. fraction)which is clearly improved compared with those recorded in a process notaccording to the invention (100 ppm by weight nitrogen) whilst stillmaking it possible to use reaction temperatures which are entirelycompatible with the duration of industrial cycles.

The examples thus show that the addition of considerable quantities ofammonia into the 2^(nd) stage reactor calms the cracking activity of theY zeolite and thus makes it possible to increase the selectivities inrespect of middle distillates. The selectivities achieved are of thesame order as those realized with silica-aluminas but with greateractivities.

The process according to the invention thus offers to the refinerconsiderable flexibility between obtaining maximized production ofnaphtha (with a low nitrogen content in the 2^(nd) stage and lowconversion in the 1^(st) stage) and maximised production of gas-oil(high nitrogen content in the 2^(nd) stage and high conversion in the1^(st) stage). This flexibility was not achieved with thesilica-aluminas used in the 2^(nd) stage.

1. A two-stage hydrocracking process of a hydrocarbon charge for the production of middle distillates comprising a first stage including catalytic hydrorefining, an intermediate separation of resultant hydrorefined products from an unconverted residue, and a second stage of hydrocracking of at least part of the unconverted residue, said second-stage operating in the presence of ammonia in a quantity corresponding to more than 150 ppm by weight nitrogen, and in the presence of a second stage catalyst containing at least one matrix, at least one Y zeolite and at least one hydro-dehydrogenating element.
 2. A process according to claim 1 in which the second-stage operates in the presence of hydrogen sulphide.
 3. A process according to claim 1 in which the quantity of nitrogen in said second stage is about 500 ppm.
 4. A process according to claim 3 in which the quantity of nitrogen is greater than 200 ppm.
 5. A process according to claim 4 in which the Y zeolite is a hydrogen-form zeolite having a SiO₂/AI₂O₃ molar ratio of 6–70, a sodium content lower than 0.15% wt, a crystalline parameter of 2.424–2.458 nm, an Na ion take-up capacity of greater than 0.85, a specific surface greater than 400 m²/g, a water vapour absorption capacity greater than 6% and a pore distribution, determined by nitrogen physisorption, of between 5 and 45% of the total porous volume of the zeolite contained in pores with a diameter of between 20×10⁻¹⁰ m and 80×10⁻¹⁰ m, and between 5 and 45% of the total porous volume of the zeolite contained in pores with a diameter greater than 80×10⁻¹⁰ m, the rest of the porous volume being contained in pores with a diameter lower than 20×10⁻¹⁰ m.
 6. A process according to claim 1 in which the catalyst of the second stage contains a matrix, at least one dealuminized Y zeolite possessing a crystalline parameter of between 2.424 nm and 2.455 nm, a global SiO₂/AI₂O₃ molar ratio greater than 8, alkaline-earth or alkali metals cations and/or rare earths cations content such as the (n ×M^(n+))/AI atomic ratio is lower than 0.8, a specific surface determined by the B.E.T method greater than 400 m²/g, and a water absorption capacity at 25° C., for a P/Po value of 0.2, greater than 6% by weight, the said catalyst also comprising at least one hydrodehydrogenating metal, and silicon deposited on the catalyst.
 7. A process according to claim 1 in which the hydrocracking or hydrorefining catalyst comprises at least one matrix, at least one element chosen from the group formed by the elements of group VIII and group VIB, and a partially amorphous Y zeolite presenting: i/a peak rate which is lower than 0.40 ii/ a crystalline fraction expressed relative to a reference Y zeolite in sodium form (Na) which is lower than about 60%.
 8. A process according to claim 1 in which the second stage catalyst contains at least one matrix doped with phosphorus, at least one acidic non-globally dealuminized Y zeolite with a crystalline parameter greater than 2.438 nm, with a global molar ratio SiO₂/AI₂O₃ lower than 8, with a framework molar ratio lower than 21 and greater than the global SiO₂/AI₂O₃ ratio.
 9. A process according to claim 1 in which the second stage catalyst also contains at least one promoter element deposited on the surface of the catalyst and chosen from the group consisting of phosphorus, boron and silicon.
 10. A process according to claim 9 in which the second stage catalyst contains as a promoter element boron and/or silicon, and optionally phosphorus.
 11. A process according to claim 1 in which the first stage hydrorefining catalyst comprises at least one matrix, at least one hydrodehydrogenating element chosen from non-noble elements of the groups VIIB and VIII, and at least one promoter agent deposited on the catalyst and chosen from phosphorus, boron and silicon.
 12. A process according to claim 11 in which the second stage catalyst contains as a promoter agent boron and/or silicon, and optionally phosphorus.
 13. A process according to claim 12 in which the catalyst also contains at least one element chosen from the group formed by the elements of groups VIIA, VIIB, VB.
 14. A process according to claim 1 in which the first stage is carried out with a conversion rate in terms of products with boiling points lower than 340° C. of between 40 and 60%.
 15. A process according to claim 1 in which an unconverted liquid residue containing hydrocarbon products with boiling points greater than 340° C. is separated by distillation.
 16. A process according to claim 1 in which the first stage of the process also comprises a hydrocracking stage carried out on a hydrocracking catalyst identical to or different from the second-stage hydrocracking catalyst.
 17. A process according to claim 1 in which, before being brought into contact with the hydrocarbon charge, the catalysts are subjected to a sulphurization treatment, and in which the first stage is carried out at 330–450° C., 5–25 MPa, with a spatial velocity of 0.1–6H⁻¹ and an H₂/charge volume ratio of 100–2000 l/l, and the second-stage proceeds at a temperature greater than 200° C., at a pressure greater than 0.1 MPa, with a spatial velocity of 0.1–20^(h-1) and H₂/charge volume ratio of 80–5000 l/l.
 18. A process according to claim 1 in which the second-stage is kept in contact with a partial H₂S pressure by adding to the charge or directly into the reactor hydrogen sulphide or at least one sulphur compound which decomposes in H₂S in the conditions of the second-stage.
 19. A process according to claim 18 in which the quantity of hydrogen sulphide added corresponds to a sulphur content by weight, relative to the charge entering the second stage greater than 20 ppm.
 20. A process according to claim 1 wherein the second stage catalyst contains a matrix of alumina gel or boehmite, a Y zeolite, MoO₃, NiO, P₂O₅ and SiO₂.
 21. A process according to claim 5 wherein the second stage catalyst contains a matrix of alumina gel or boehmite, a Y zeolite, MoO₃, NiO, P₂O₅ and SiO₂.
 22. A process according to claim 6 wherein the second stage catalyst contains a matrix of alumina gel or boehmite, a Y zeolite, MoO₃, NiO, P₂O₅ and SiO₂.
 23. A process according to claim 7 wherein the second stage catalyst contains a matrix of alumina gel or boehmite, a Y zeolite, MoO₃, NiO, P₂O₅ and SiO₂.
 24. A process according to claim 8 wherein the second stage catalyst contains a matrix of alumina gel or boehmite, a Y zeolite, MoO₃, NiO, P₂O₅ and SiO₂. 